Hydrocracking of gas oils



- May 5, 1959 K. K. KEARBY- Em 2,885,346

HYDROCRACKING OF G AS OILS Filed March 17. 1953 8 Sheets-Sheet l ADSORBER in as LHYDROGEN FEED 37 FRACTIONATOR 4 HYDROCRACKI NG R EACTOR GAS-LIQUID SEPARATOR FIGURE-I SEPARATOR FEED Kennefh K. Kear by Isidor Kirshenbaum By 7 m Attorney Inventors May 5, 1959 Filed March 17, 1953 K. K. KE ARBY ET AL HYDROCRACKING OF GAS OILS 8 Sheets-Sheet 2 E 2 a i v 3 i 2 g T 5' g 5; (DEC 2 :0 go 0 ad- I: 0 $5 i a: A

In I Q I f 5 p- E T Q P. Q 3 5 H m 2 U. u 9 2 1 OE: 0 m r.- 9 I 9 o 5 n:

m Eu I 5 5 z L -Ll p (D m a 9 g I 3 9. l Q h 1 2 Kenneth K. Kearb y Inventors Isidor Kirshenbaum By )7, 214 Attorney FIGURE "2 May 5, 1959 K. K. KEARBYY ETAL 2,885,346

HYDROCRACKING 0F GAS oILs Filed March 17, 1953 8 Sheets-Sheet 3 LEAD UNIT 300 HYDROCRACKING REACTOR TAIL UNIT HYDROCRACKING REACTOR FIGURE-3 Kenneth K. Kearby Isidor Kirshenboum Inventors By 1% Wmforney Filed March 17, 1953 M y 5 K. KEARBY ET AL 2,885,346

HYDROCRACKING OF GAS OILS 8 Sheets-Sheet 4 HYDROCRACKING REACTOR Kennefh Kea rby Isidor Kirshenbaum MVEITTOFS By J, WAfforney y 1 K. K -KEARBY ET AL 2,885,346

" HYD-ROCRACKING OF GAS OILS Filed March 1'7. 1953 8 Sheets-Sheet 6 CARBON FORMATION AT 9 T 28% GASOLINE YIELD CARBON-BASED ON FEED 01 o 5 so I5 20 25 WT. *7, Mo 0 ON CATALYST FIGURE-6 WT. CONVERSION FEED RATE V/V/HR.

0.2 013 64 0:5 ofs of? FEED RATE V/V/HR.

FIGURE-7 I v FIGURE-8 Kenneth K. Kearby v lsidor Kirshenbaum By M 7Aw Afforney Inventors K. K. KEARBY ET AL HYDROCRACKING OF GAS OILS May 5, 1959 8 Sheets-Sheet 7 Filed March 17, 1953 HYDROCRACKING OF E.T.L.G.O.

5 nw 9 w 3 w W m m E O U L m o o I 5 mm 0 M. G w 1. W w m 4 w P 5 4 m .C O 6 5 o 5 w w a W 3 3 2 w m m V m 0 8 m w 0 G VK m 6 0 m e O 0 K 43 4 O 2 O Isidor K irshenbaum y gar 71W! Attorney yS, 1959 K. K. KEARBY ET AL 2,885,346

HYDROCRACKING OF GAS OILS Filed Marqh 17, 1953 8' Sheets-Sheet 8 FIGURE 'lI HYDROCRACKING CRACKING E |ooo|o2or-? O psig S 945 F/ g 200 psig HYDROCRAGKING IOOOF. 2o 100-500 psig 30 4o 50 e0 70 so 430 CON.,VOL."/.,

Kenneth K. Kearby lsidor Kirshenbaum Inventors 1 6 5 4 H Q Q ACKING 05F GA OILS t 7 Kenneth K. Kearby, Cranford; and'lsidor Kirshenbauln, Union, N J assignors to Esso Research and Engineering Company, a corporation of Delaware I V Application .March 17, 1 953,!Serial No. 343,919 7 Claims. (Cl. 208-79) The .present'invention relates to a new and improved 'hydrocracking process for'hydrocarbons suchas gas oils and other petroleum fractionsuheavier .than naphtha.

" This -process is designed to obtain high qualityfuels of gasoline: boiling range. 1 The invention relates furtherto .a process which secures good yields of high quality fuels accompanied by reduced. production-of coke andcertain other. degradation products .as comparedwith more con- .ventional refiningoperations. The invention isparticularly applicable to hydrocarbon feed stocks of the general .type mentioned above having high contents of naphthenes. It. is. not necessarily limited to the highly naphthenic stocks, however. U .Gasoline, as theterm is commonlyusedtodefinemotor fuels, is commonly produced from petroleum by the following principal processes:

t (1), By simpledistillation. or which produces a .virgin? gasoline. usually of rather highly saturatedtype,

.lyapplied to naphthasandlighter fractions, such as alkyl- 'ation, isomerization, polymerization andv aromatization,

to obtain motor fuels fof highlyibranched parafiinic type. These are frequently ofrelatively. .very highoctane value.

Commercial'motor fuels are usually obtainedby blending together suitable proportions of two. or more of. the above products, adding special modifiers such as tetraethyl'lead, etc., where desirable. V Of the foregoing,the catalytic cracking processes have received by far thegreatest emphasis in recent years. By subjecting gas oil fractions of the 400 to 800 F. boiling range, for example, to catalytic cracking. operations, motor fuel constituentsof'very good quality .(high octane .number) may be produced at moderate cost.., Some of the special synthesizing or reforming techniques mentioned above are used to produce particular ingredients for special requirements (e.g. constituents foraviation ,fuels)-:but catalytic cracking has been used more widely for 'general'fuels because of its efiiciencyand economy. 'Under some conditions, however, catalytic cracking is not a preferable process. Modern cracking systemsop- 'erating at or near atmospheric pressures require large and expensive-apparatusand they tend to produce rather large proportions of coke and otherndegradation products. Moreover,'cracking catalysts are easilypoisoned and rendered ineffective by some commonlyocc'urring elements. This requires constant or frequent catalyst replacement, resulting inhigh catalyst costs. Thus, with certain types of oils containing substantial proportions of certain types.

and final boiling points may be I present invention to high'quality; .At the same time the process ofthisin'ven- -tionl is accompanied .by

tiabproportions of cracking treatment in the presence of aspecial type of "catalyst I described octane number is obtained;

from 8 to 16% 2,885,346 Patented May...5, 19.59

catalyst poisoning constituents, cracking catalysts are rapidly rendered ineffective. Also, catalytic crackingis not as effective in upgrading certain kinds of feed stocks,

:for example someof those of high naphthenes content,

=or alumina type mightbe modified by adding thereto hydrogenation or dehydrogenation catalysts of-known Other prior art has suggested modifying alumina or-silica catalysts by mixing or impregnating them: with various salts and other catalytically active materials' including, for example, the chlorides, oxides, sulfides, etc., or. molybdenum and of various other metals. Theme of an improved catalyst. of this general type is one feature of the present invention.

.- In particular, however, the invention is based on the discovery that certain petroleum fractions 'which: have been diflicult to convert efliciently to motor fuels may be more effectively converted than hitherto by a careful and particular choice of operating conditions, catalyst, etc.

For example, a gas oil fraction of 400 to 800 F. initial treated according to the produce motor fuels of exceptionally the production of remarkably small quantities of'carbon'or. coke and by low production ofcertain' other light end products, especially of C; and C hydrocarbons. For example, a gasoil feed having the boiling range mentioned above and containing substannaphthenes may be subjected to hydrobelow. Moderate pressures," preferably between and 350 p.s.i.g. may be used and the gas oil fraction is converted in the presence of substantial proportions of hydrogen gaswhich is passed through the process at the rate of 3000 to 10,000 cubic feet (a't standardatmospheric pressure and temperature) for each barrel of feed. i I Y I From this process'a good yield of motor fuel ofhigh quite contrary to experience with somewhat related prior art process Normally a gas oil -product,'hydrogenated under pressure, would be expected to produce motor fuel of'substantially lower octane number than a regular catalytic cracking'process.

In other words, hydrogenation under pressure would be expected from experience in the prior art to produce a substantially saturated and relatively unbranched fuel product. I iii H A:catalyst' which is particularly effective-and which ispresently preferred for the present invention consists of a, conventional silica-alumina cracking catalystypreferably predominating 'in silica, e.g. an 87% silica- 13% alumina catalyst-of well-known type, to Which'isadded preferably around 10 to 12% by weight, based on the total catalyst composition, of molybdenum -When such a combination catalyst is used in the process described below to simultaneouslycrack and hydrogenate or tohydrocrack a light gas oil, not only isa motor fuel of particularly high octane number obtained but it isrobtained in very good yeilds. At the same time a very small amount of carbon or coke is produced; The amount of coke may be less than 1% of-thefeed when fluid-solids technique is employed. With other conversion procedures, coke may reach about 2% but even this is less than might be expected. a

The essence of the present invention, therefore, is the discovery that a higher octane gasoline can be produced with minimum coke formation in a single cracking-hydroforming operation, i.e. by cracking with a combination catalyst in the presence of hydrogen and under moderate pressure. The temperature range employed in this process is somewhat higher than in most of the prior art, i.e. about 970 to 1100 F., preferably between 1000 and 1070 F.

Catalysts suitable for this invention are satisfactorily prepared from conventional silica-alumina cracking catalysts or from catalysts of this general type but of reduced cracking activity. Thus, they may be prepared from used cracking silica-alumina catalysts which have been already subjected to conventional catalytic cracking operations and have lost some of their activity. Much less satisfactory catalysts are prepared from conventional cracking catalysts which contain MgO and SiO, as their active components. The silica-alumina type catalysts are preferably impregnated with a conventional hydrogenating component, preferably molybdenum oxide. Catalysts of the type required for the present invention are somewhat costly but fortunately their costly molybdenum content may be reclaimed by volatilizing molybdena from the used catalyst and depositing this ingredient either on a fresh base or on used cracking catalyst of the desired range of activity. The used cracking catalyst, when it is chosen, is preferably of the specific type mentioned above and widely employed in catalytic cracking of gas oils. It consists of about 87% SiO and 13% A1 by weight. The hydrocracking catalyst product prepared therefrom by addition of molybdena is preferably regenerated after use in a conventional manner. Air is used to burn the coke off the catalyst and also to oxidize any molybdenum or molybdenum oxides, etc., adhering thereto. If additional heat is needed to carry out the regeneration a tail gas from a conventional catalytic cracking reactor may be introduced and burned with the coke on the catalyst, a suitable oxidation gas, usually air, being introduced in appropriate proportions. The specific manner preferred for reactivating the catalysts or impregnating them with molybdena is described in greater detail below but briefly the molybdenum oxide or equivalent is led into the regenerator where a stream of the catalyst is being regenerated and is added to the base catalyst there. The invention will be more fully understood by reference to specific embodiments which are illustrated graphically in the attached drawings.

Figure 1 is a diagrammatic view of a cracking and dehydrogenating or hydrocracking system embodying the present invention;

Figure 2 illustrates diagrammatically another system embodying the general features of the invention;

Figure 3 shows a multiple reactor hydrocracking process according to the present invention which results in high yields of high octane number gasoline with a minimum degradation to carbon and fuel gas;

Figure 4 shows diagrammatically a hydrocracking system including a unit for activating or reactivating conventionally used cracking catalyst to make it suitable for purposes of the present invention;

Figure 5 shows another modified system;

Figure 6 is a graph showing the relation of carbon formation to the hydrogenating catalyst component of the hydrocracking catalyst;

Figure 7 shows graphically the relation between feed rate and selectivity to carbon in a hydrocracking process carried out at 1,000 F. under a pressure of 100 p.s.i.g.;

Figure 8 shows the relation between the feed rate and the percentage of conversion of a gas oil to high grade gasoline under the same hydrocracking conditions as Figure 7;

Figures 9 and 10 show graphically the relation between total conversion and yields of C --430 F. and C -43O F. gasoline;

Figure 11 shows graphically the effect of process variables and conversion levels on the olefinicity of products.

Before proceeding to a detailed description of the various systems and modifications shown in the drawings, it may be noted that carbon formation apparently can be minimized during cracking by (1) decreasing the transfer of hydrogen and minimizing dehydrogenation reactions of the coke precursors; (2) hydrogenation of coke precursors, and (3) decreasing catalyst activity to permit the use of higher cracking temperatures. All of these objectives are obtained in some degree by hydrocracking over a catalyst of the type described above, particularly an alumina or silica-alumina catalyst carrying the optimum percentage of dehydrogenating catalyst, the latter preferably being M00 The M00 appears to decrease hydrogen transfer activity and in the presence of relatively high pressures of hydrogen appears to hydrogenate the coke or carbon precursors. As a result of these facts it appears that at least 30 to 40% less carbon may be obtained in hydrocracking as compared with conventional catalytic cracking to obtain the same yield of 0 gasoline. The process also produces a gasoline containing more olefins and less aromatics.

In addition to producing less carbon the process of the present invention produces significantly less proportions of C and C gases. However, the process tends to produce more dry gas (C to C than does catalytic cracking under similar conditions. This tendency to produce more of the dry gases and less 0 and C fractions may be reversed by treating the catalyst with small amounts of H 0 or HF and this is an additional feature of the invention. The process therefore is flexible and can be adapted to situations where production of the lighter hydrocarbons (C and below) is desirable, e.g. where liquified or high pressure gas is in good demand.

A catalyst composed of about 87 parts by weight SiO and 13 parts A1 0 (a silica-alumina catalyst) and 10 parts M00 Was used in most of the experiments described below. However, the M00 or equivalent dehydrogenation component may be varied somewhat in most cases. It may be added by merely mixing the dry ingredients together or by preparing a solution of ammonium molybdate, impregnating the silica-alumina catalyst with this solution, and drying and calcining. Alternatively the M00 may be sublimed to coat the cracking catalyst. Other methods known to those skilled in the art may also be used.

In broader respects the silica-alumina component usually should comprise from 84 to 92 parts by weight of the total catalyst combined with 8 to 16 parts of the hydrogenation catalyst. The latter may be promoted, if desired, by treatment with zinc oxide, phosphoric acid or phosphates, sulfides, calcium oxide, cobalt oxide, magnesium oxide, conventional halogen compounds, or even by small quantites of potassium oxide. Such promotion is well known in the art. The cracking component of the catalyst also may be promoted with phosphoric acid, boria, hydrogen fluoride or other fluorides or other halogens. For high sulfur feed stock special advantages may be obtained by using equivalent quantities of cobalt-molybdate or zinc-molybdate as the hydrogenation catalyst in place of the molybdena.

The process of this invention is primarily designed for cracking gas oils but it also may be used to advantage in cracking other fractions such as petroleum residues, waxes, cycle stocks, kerosenes and other middle distillates. Synthetic feed stocks of appropriate grade may be used such as those formed by hydrocarbon synthesis as well as the products of visbreaking and catalytic polyforming. The process may be used also for liquid phase as well as vapor phase hydrocracking, although it 1s applicable primarily to the latter. The process is useful in conjunction with hy-droforming systems where an excess of hydrogen is produced. It may thus be used to obtain a desirable balance in refinery operations that could not be secured with conventional cracking processes and apparatus.

5 in generalg 'it is"preferred-amuse the fluidssolids 'tech 'nique for hydrocrackin'g, treating the feedstock 'in the vapor phase-asis -well knowrr in the catalytic cracking art-The invention will be described in detail-particularly. .withrespect to the fluid catalyst process but it will -:be understood that fixed bedand moving bed catalysts ofaother conventional--typ'es-'maybeeinployedr t In preparing 1 the hydrocracking catalysts vused in the present inventio'n various'methods may. be used'. Moldbdena may beadded as molybdena blue, bycarbonyl decomposition; or'it may.- be- -added the naphthenate or asrammonium molybdate. 'With some feed stocks and in some:types of operations it' iswadvantageous'to prepare the aluminum "oxide component of the catalyst in carbonicfiacidt-solutions.-= 1 f In operating the hydrocracking- -proeess, improved selectivity'to 'high octane number zfuel and/or to C and C compounds may beobtained by adding small amounts of acids such as HFyHGlyacetic acid,- naphthenic acids, etc. *J-The process, therefore, is of special advantage for producing aviation-fuel. Y

Hydrocracking catalysts appear to -be less susceptible to-loss f activity and selectivity caused by high aromatic feed stocks thaniare conventional cracking catalysts: in some cases the-hydrocracking catalysts used in the present invention appear to be benefited by recycling aromatic streams or by pretreating the catalyst with aromatics such as benzene and toluene. 7 7

Referringnow to Fig. 1 a system is disclosed wherein 'a-suitable feed stock such as a gas oil, preferably a virgin gas oil, is fed through line 11 to a conventional cracking unit=13.'---The products--from the cracker are passed through line 15 .into a 'gas liquid separator 17. Gaseous products arelremoved from a separator through line 19 and the liquidproducts are takenthrough "line 21 into a fractionator 23; Thecycleoilfraction from the fractionator, which maybe all'orpart of the 430 F.+ product is' -then led fromthe bottom of the fractionator through line 25 to a hydrocracking unit 27. Valves 29 'and31' may be .provided for withdrawing a portion of this product through line 33 if desired.

Inthe hydrocracking 'unit theoil is subjected to a temperature 015750 to- 1100 F.,- preferably between 1035 and 1070 F. 'Ihe hydrocrackingoperation is carried out at 100 to 500 p.s.i. pressure, preferably about 150- 350 psig. -Most prior art processes have required higher pressures.- Hydrogen rich gas is-suppliedthrough line 35 atthe rate of about 500 to 10,000 cubic feet of hydrogen,

preferably at least 3000 cubic'fee't, on an atmospheric pressure basis, per barrel of oil fed to the hydrocraking unit. *The catalyst preferred contains about" 10% M00 on a silica-alumina (87-13) base as described above. The hydrogen may be supplied from external sourcesor may be recycled as'willbe pointed out below.

The gaseous products "from the hydrocracking unit are taken overhead through a line 37 to :a gas-liquid 'separator 39. Dry gases'from the separator -are-removed through line 41 'andmay be recycled through'line 43 or-*drawn off through valve 45. The liquid products from the separator are withdrawn through line 47 and passed into fractionator 23. Although shown as one fractionating'unit' a number of fractionators may be used 'and -the'products from the cracking reactor 13 and the hydrocracking reactor 27 may be led to separate fractionators'ifdesired."

The dry gases from separator 39 contain considerable hydrogen and arepartially orcompletely recycled to the hydrocracking unit, either through lines- 49 and 35 or through lines 51, adsorber 53 and lines 55 and'35. As is well known; an'adsorberserves to concentrate the hydrogen content of the recycled gas. The adsorber unit may utilize the fluid charadsorption process or may use oil adsorption or low temperature adsorption processes. Appropriate valves 57 and59 are-provided to control the flow or to bypass the adsorber as desired. Addi- 6 tional 'hydrogen is sometimes needed from'fextraneous sources and such may be introduced through line' '61 by opening valve 63. Extraneous hydrogen-may also be introduced through line 65 ahead *of-thea'dsorber to improbe adsorber operation under some conditions.-

1" A convenient source of extraneous make up hydrogen is the tail gas from a hydroforming unit of-known type. The tail gas from thecracking unit- 13 may be used, however, after enrichment by a suitable adsorption process. With some feed stocks no enrichment "of'hydrogen may be required. In other cases the makeup hydrogen can be manufactured in a converter from natural gas or other conventional sources. Under some conditions it may be desirable to introduce fresh feed to the hydrocracker along with the cycle oil from line 25. Inthis case the fresh feed-maybe'in'troduced from its supply through'line 67 connected to line lland controlled by valves 69 and 71: In one modification of the invention all or partof' the fresh feed is passed over a hydrocracking catalyst 'in' the presence of hydrogen at-atemperature'inthe' range :of 500 to 750 F. before passing the feed to the cracker. This conditions the feed by-removing crackingcatalyst poisons such as nickel, vanadium, iron, sulfur, etc: It will (be understood that gases other than hydrogen produced by the hydroeracker may be suitably utilized in various ways. The C to C olefins contained therein can be concentrated or passed directly to a polymerization unit. This reduces or eliminates compression coststas compared with gases from an atmospheric pressure fluid crackingunit. Refer-ring now to Fig. 2; freshfeed,;for'example,one which contains a substantial proportion er naphthenic stock, is introduced through incoming 'feedtline101 into a hydrocracking unit 103 operating under the "same general conditions of temperature, 'hyd'rogempressure and catalyst .as that shown in Fig. ln'The productof'the hydrocracker is removed overheadthrough aline 105 to a fractionator-1-07. It may be fractionated via knockyout dnums and fractionation columns. Part or all of the gas product from thefractionator ma 'be'recycled: through lines 108, 109, and 111 or alternatively through lines 108, 113 and adsorber 115 for-hydrogen enrichment as. in Fig. 1. Part of the product gases maybe'withdrawn from the system through lines 117 and 119, suitable valves 121, 123 and 125 being provided. 1 i i Ordinarily the hydrocracker is operatedunderconditions of low overhead or netgas production so thatno net hydrogen is consumed." Under moresevere'operating conditions, as in Fig. l, make-up hydrogen may be required and this can beadded through line 127i"; rlternatively the tailgas from cracking-unit to be' des'cri bed may also be recycled through line 129: A hydrogen donor may be utilized but this forms no part of the persent invention. 1 The 430 'F.+ stream from the bottom 'of fractionator 107 is led through'flines 131, 133, into a cracking unit 137 which is of conventionaltype; preferably utilizing the fluidized 'solids technique. Herejthe feed is cracked, and the cracked products areledvialine 139 to fractionator 141,- a knock-out system lbeingincluded in line 139 if desired, From the, fractionatorcycle oil.is sent back to the cracking. unit-via.lines.143-, 14 5 and,;135. Fractionators 107 and.141 maybe:- combined in- .a.;single unit ifdesired. Regeneratorsfor cracking and/or hydrocracking catalysts may be usedbut arenot shown The aromatic and olefinic constituents of thefin prod; ucts maybe controlledrby recycling part of. theproducts from. the fractionator 141 (or 107) through-lines.-147,\149 and 129 tothe: hydrocracking unit.- In some cases 'it' is convenient to recyclesome of the heavy fractionsof the cycle oil back to thehydrocracker and this may be done through lines 143, 151 and 129. The hydrocracker isable to crack these streams with less carbon formation than is -obtained in conventional cracking. This is an important fresh feed directly into the cracking unit. Line 153, controlled by valves 155 and 157, is provided for this purpose. Heavy bottoms from the fractionator 141 can also 'be Withdrawn through line 159 under the control of valves 161, 163 and 165. Likewise the products of line 147 may be withdrawn from the system through line 167 controlled by valves 169 and 171 instead of being recycled.

The combination cracking and hydrocracking process of Fig. 2 is especially applicable to high nitrogen and high sulfur feed stocks. The hydrocracking step lowers their nitrogen and sulfur contents. This permits the use of a sulfur sensitive catalyst in the cracking unit.

The hydrocracking operation of Fig. 2 is carried out preferably at 150-350 p.s.i.g. and at l035-l060 F. using 3,000 to 10,000 cu. ft. of hydrogen gas recycle per barrel of oil. When hydrogen enrichment is carried out by the fluid char adsorption process, the adsorption is carried out in the pressure range of 100-500 p.s.i.g. at a temperature of ISO-200 F. A char circulation rate of 0.25.l lb./ cu. ft. of tail gas is most satisfactory.

In some cases feed stocks contain compounds that are difiicult to crack and do not give the best operation in the units or systems described above. If the hydrocracking process is operated under conditions severe enough to crack these refractory compounds some of those compounds which are cracked more easily break down to gas with excessive coke formation. According to the present invention this difficulty may "be overcome by the use of a multiple hydrocracking reacting system in which the lead reactor or plurality of reactors are operated under mild conditions and the subsequent reactors under severe conditions. Such a system is illustrated in Fig. 3.

In the system of Fig. 3 the product from a lead or first reactor is separated (or the separator may be bypassed) and the cycle oil boiling about 430 F. is sent on to a subsequent or tail reactor. Thus, the feed enters reactor unit 300 through a line 302 along with hydrogen gas which is recycled through a line 304. The product of this first reactor is carried overhead via line 306 to a knock-out drum and/or a fractionating column indicated diagrammatically at 308. The recycled fraction is then led through lines 310 and 312 to the second or tail reactor 314 along with recycled gas from the fractionator fed through lines 316 and 318. In some cases it is advantageous to add some fresh feed to the second reactor through a line 320, valves 322 and 324 being provided for proprotioning the feed to the separate reactors. The products from tail reactor 314 are taken overhead through line 326 into line 306 and into the knock-out drum and fractionator. Obviously separate drums and/ or fractionators may be used for the separate reactors if desired and this is often of advantage when the entire process is operated at substantially one pressure. However, if the two reactor units are operated at different pressures, as is sometimes desirable, it may be preferable to keep the liquid gas separating equipment separate for the difierent stages. This minimizes the need for compressing the recycled gas in one case or the other.

The hydrocracking reactor units may be operated so as to either produce or consume net hydrogen. Where net hydrogen is consumed in the operation it is necessary to supply make-up hydrogen. This may be obtained either from other reactor units operating under different conditions or from an extraneous source, such as a hydroformer.

The lead reactor units are operated under conditions previously specified, i.e. preferably at 100 to 500 p.s.i.g. or more specifically 100 to 350 p.s.i.g. and at a temperature ranging between about 1000 and 1060 F. with 3000 to 10,000 cubic feet of hydrogen gas per barrel of oil.

The tail reactors may be operated at much higher pressures in some cases, for example at pressures up to 5,000 p.s.i.g. and at temperatures of 970 to 1100 F. A hydrogen recycle rate between 1,000 to 15,000 cubic feet per barrel of oil may be used in extreme cases. As in the case of the lead reactors various hydrocracking catalysts may be used. It is desirable, however, to activate the catalysts for the tail reactor by addition of halogen, for example by treatment with HF, NH F or HCl. In addition to the hydrocracking catalysts mentioned above excellent catalysts for the second stage are, for example, silica alumina compositions, e.g. of '87SioO /13Al O with or without halide promotion and without molybdenum oxide, catalysts based on HF treated clays with HF promotion, etc. When catalysts of the same composition are used both in the lead and the tail reactors a common regenerator, not shown in Fig. 4, may be used. Otherwise separate regenerators may be required.

As previously described, it may be desirable to recover the molybdenum oxide from spent catalyst since losses of this expensive catalyst ingredient frequently are considerable. For example, in a 20,000 barrel per day fluid cracking catalyst plant it is common practice to reject about 1.5 tons of catalyst each day. This amounts to a loss of over 45 tons of molybdenum oxide per year. According to the present invention this molybdena may be recovered in situ by leading off a side stream of coked catalyst to a separate vessel where the molybdena is evaporated and stripped ofi at a temperature of 1300 to 2,000 F. A system for carrying out a hydrocracking operation and recovering the expensive molybdena ingredient is illustrated in Fig. 4.

As shown in this figure hydrocracking is carried out in a reactor 400 from which used catalyst is removed through a line 402. This catalyst is fluidized and introduced into a used catalyst zone 404 of a catalyst preparation unit 406 through lines 408 and 410. Gas from the regenerator 412 is supplied through lines 414 and 416 for fluidization. The advantage of using this gas from the regenerator is that it normally contains appreciable quantities of M00;, lost from the catalyst during regeneration as well as containing some steam. This steam aids in stripping the molybdena from the used catalyst.

The temperature in zone 404 is maintained between 1300 and 2,000 F., preferably 1400 to 1600 F., by introducing excess air via lines 418 and 420. This air is used in burning the coke off the catalyst as well as in oxidizing the molybdena. If additional heat is needed tail gas may be drawn from the reactor through lines 422, 424 and 420 (via a stripper or fractionator described below). Auxiliary fuel such as gas, torch oil, or coke may be used if needed. This gas or other auxiliary fuel is then burned with the coke on the catalyst in the catalyst preparation unit.

The volatilized molybdena is carried from the zone 404 into an upper zone 430 of the preparation unit 406. The two zones are separated by a grid 431. Here the molybdena is deposited on a new catalyst base which is introduced through line 432 from any suitable source. The renewed catalyst is then removed via line 434 and 436, through which it is returned to the reactor 400 for use.

If pretreatment of the catalyst, such as calcination, for example, is desired this may be carried out in situ in the catalyst preparation unit 404, 430 by introducing hydroformer tail gas or air via line 420. Alternatively, the prepared catalyst may be transferred to separate vessels for the calcination or pretreatment step. Make-up molybdena may be added to zone 404 through a line 438. Rejected base catalyst may be withdrawn through line 440. It will be understood that the catalytic solids are preferably fluidized and appropriate fiuidizing means, draw-off means, standpipes, grids and distributors, etc., will be used as needed. These are well known in the art and need not be described in detail.

In a modification of this invention, Fig. 4, the gas stream :richinimolybdenamay be led directlytback; to :the -r.eactor -or-;regenerator..by sQ idirecting line 442 to con- .nect. for example, with aline 444. :In-this case a stream tofm-ake-up .base catalyst-would be added to the reactor -or regenerator. .Regenerated catalyst from the regenerator. is returned'to the reactor through line 444.. Hydrogen rich gas may be recycled. from-the stripper or fractionator 446 through a line 447, 448, .appropriate valves 3450.:and 452-being provided for directing the gas from the latterunit as desired..-- Y

- Othermodificationswithin the scope of the invention are-the use of catalyst preparation zones, the use of either purifiedor commercial grade molybdena to:make .up .for losses and the design of the molybdena recovery lsection as. either a fluid roaster rotary kiln or a Herre- .-shoif furnace, etc. The volatilizing gas for'recovering molybdena may be any inert gas and more molybdena may be added if needed.

---=-As pointed out above, there-area number'of advantages in submitting high boiling distillates to hydro- :cracking over acatalyst which comprises a cracking catalyst base carrying an active'molybdena component. One major advantage-is the marked reduction in coke formation. Y For'example, hydrocrackinga light gas oil overa catalyst having acompositionof 90% (87SiO 13Al O 10% M composition actually-produceda 99 research octane number gasoline with only 0.9% by weight of carbon based on the feed. Acon- -ventionalcracking process using the same feed stock over the base catalyst with no molybdena content to obtain the same gasoline yield and octane number produced 3.9% of carbon by weight. Catalysts of this type :areuseful not only in catalyticcracking of gas oils but also in catalytic reforming of naphthas. -=---Although catalystsof-thetype just "described have excellent selectivity for motor fuel productionwhenpre- 'paredfrom fresh or steam treated cracking catalyst-base "there are a number of economic advantages in preparing the catalyst either for hydrocracking or for naphtha reforming, from used cracking catalyst. Moreover, with some feed stocks this method of preparation results in abetter-distribution of products; --The hydrocracking catalyst made from used cracking catalyst can be prepared by conventional methodsas described above.- Howeven'where refinerieshave cracking and naphtha reforming reactor units, especially those of the fluid solids catalyst type, a" preferred method "of catalystpreparation is shown schematically in-Fig;

The catalyst preparation-unitgin this modification may consist of one vessel 500 have two zones 502and 504. In;thelower Zone molybdena is fluidized while in "the upper zone a cracking catalyst. is fluidized. Freshmolybdena may be fedin through line .506 "and fresh b'ase catalyst through line 508 when desired. Fluidizingadistributing, draw-off ,rneans, etc., may beprovided as..is well known; 'In general,. however, regenerated used cracking catalyst is led .from a ,regenerator unit 510 through lines 512 and 514 into zone 504 of the catalyst preparation unit. This unit is maintained preferably at a temperature of between 1,000 to 1300 F. bythe hot freshly regenerated catalyst. Temperatures as high..as -l 5 00 F. are permissiblein this zone. Heat may bejcontrolled in zone 504 ,by recirculating"'unregenerated"or partially regenerated catalyst via lines 515,"516"and1514 directlyfrom catalytic cracking reactor 518. This'unregenerated catalyst, coated to some extent with cok'e, may be burnedin .situto. supply the additional heat which may be required under some conditions. Air for burning such coke or carbon may be-introduced either through ithejibasej catalyst. supply.line.508 or through ai linet520 leading from: a hydrocracking or reforming unit shown atztheright-of Fig. 5.. 1. The lower fluidizedmolybdenazone 502 is maintained preferably at a temperature of 1400 to 1600 F. al-

though lower or higher temperatures may be used. The

*selectivityuto carbon as the feed-rate is increased.--*Fig. 381 shows "a corresponding conversion (catalyst activity) =overhead from a hydrocracking regenerator unit as shown diagrammatically at 530. These gases pass througha :line 532% under'control of. valves 534 and 536-.into line 538 and thence into line 520 previously'mentioned: Such gases contain MoO ..which.is normally 'lost duringwre/ generation. -They also-contain some steam. --The steam aids in volatilizing the molybdena in zone 502. Addi- -tional'steam, air,or=inert gas may be added via lines 540,542 and 544, as. desired. Itis often convenient-to add a tail gas from the cracking or hydrocracking unit: together with some air'from dines 540; 520. .This mixture of tail gas and air burns l-Zlll' zoner502andthus helps to maintain the desired high temperature of the molybdena. Y lt will be understood that the hydrocracking reactor 550 is operatedsuch thatcatalyst is passed through line .1552 to the regenerator 530 and returned through "lines 554, 556. Newly prepared catalyst'fromthe preparation unit 500 may be-fed into the hydrocracking line 556 through a line 558 as will 'beobvious. :Theopraition of the zicrackingreactor and regenerator 518 and 510 is conventional and need not be further des'cribedrw; Q As'indicated above, an important feature of the present invention; is 'thediscovery that optimum operating conditions may be obtained by a careful selection and controlof catalysts as well as of feed rates, temperatures 'andspressures; 2 Figs 6 shows theremarkable reduction in carbon-formation at a 28%' gasoline yield when cracking gas oil of a conventional quality. As will be noted, the carbon content based on the feed was over 12% when a straight cracking catalyst was used and this dropped to a very low level near 2%, when about 8 .to 10%.wby weight o'f molybdena-was added to thecatalyst. As the molybdena content exceeded about 16% there 'was an appreciable risein the carbon formation, the C to-460": F..,gasoline-yieldbeing kept constant.

.'As shown inv Fig. 7 theprocess shows a decreasing decreaseas the feed rate is increased. Hence, the'factors of? carbon formation and conversion must be balanced to obtain optimum operating conditions. The feed rate selected will vary with feed stock, with catalyst, and with operating conditionsf For the feed stock, catalyst'and'conditionsshown above, the optimum appears to'be between"about0.5 to 1.0 v./v./hr.

As stated above',"'an optimum temperature should be selected so thatneither excessive gas nor carbon is produced 'andso that gasoline .of desired octane value is obtained. ,Too low temperatures of operation give low .octanesntunber .gasolines, 10w olefinicity in gasoline and more; carbon. Data-are shownin the table below and in FiguresS, 10, and 11. Preferably the'temperature is between 1035". -and-l070 F. The preferred-pressure isabetwcen 15,0 andz350; p.s.i.g., about 200 p.s.i.g. being desirable; Thep-referredz hydrogen recycle rate is at least.,4,000.. cubic feet per barrel of oil, with about 10,000 asamaximum. 1

A -comp'ariso'n was made between two catalyst bases (at) 6nliof 65%=-Si0g'and 35% MgO and (b) one of 87% SiO; and 13%Al O Each of these was treated to add thereto 10% by weight of molybdenum oxide.

Both were used to hydrocrack a light gas oil of 490 to 700 F. boiling range. The pressure used was 200 p.s.i.g., temperature 1000" F., and a hydrogen feed rate of 3000 cubic feet (standard conditions) per barrel of It was found that at 50% conversion level, an increase in operating temperature from 1000 to 1050 F. raised the octane number from an average of 90 to an average of 92, other conditions being as indicated above.

Hydrocracking of gas oils over the catalyst of Table I produces about 1-2% more C gasoline and about 25% more C gasoline than does catalytic cracking to the same conversion level at 10001020 F. in the same unit at atmospheric pressure. However, the hydrocracking process produces about half as much C 's and C s and about 3% more dry gas than does high temperature catalytic cracking. At about a 50% conversion level, the hydrocracking C 430 F. gasoline had a 93 octane number as compared to 95 for the catalytically cracked gasoline.

A further study was made to compare hydrocracking of an East Texas light gas oil (490-700 F. boiling range) with the same catalyst and hydrogen feed as used for the data of Table I, a temperature of 1000 to 1020 F. Pressures of 200 p.s.i.g. were used for hydrocracking and atmospheric pressure for conventional catalytic cracking. Results are shown in Table II.

When hydrocracking and cracking were carried out, as in Table II to obtain identical yields of 0 gasoline (19%), the hydrocracking produced 1% or less of carbon whereas catalytic cracking produced 1.4%-an increase of at least 40%.

Referring to Fig. 11, a marked increase in olefinicity was obtained by hydrocracking as compared with ordinary cracking with standard 87 silica-13 alumina type catalyst. Within conversion ranges of 40 to 60%, the bromine number of the cracked product (1000-1020 F.) ranged between about 40 and 65 as compared with about 50 to 75 for the hydrocracking product at the same temperature at 200 p.s.i.g. Other hydrocracking pressures as low as 100 and as high as 500 p.s.i.g. fell substantially on the same curve. With cracking under pressure of 200 p.s.i.g. the temperature had to be reduced to 945 to keep the conversion level below 50%. A bromine number of about 39 was obtained (point X) suggesting that cracking under pressure may tend to reduce olefinicity. Upon increasing hydrocracking temperature to 1040 F. (at 200 p.s.i.g.) the bromine number was increased considerably further in a series of runs over various conversion ranges. Further increases in temperature to 1050 and 1070 F. did not appreciably change the bromine numbers, the data falling substantially on the same upper curve.

What is claimed is:

1. A process of cracking mineral oil fractions boiling above the gasoline range to produce motor fuel of high octane number with low carbon production which comprises subjecting said fractions to contact with a cracking catalyst containing a major proportion of silica, a minor proportion of a refractory metal oxide selected from the group consistng of alumina and magnesia and from 8 to 16% of molybdenum oxide at a temperature of 1000 to 1100 F. and a pressure of 150 to 300 pounds pressure per square inch and in the presence of hydrogen gas supplied at the rate of 3000 to 10,000 cubic feet under standard conditions per barrel of feed.

2. A process of cracking mineral oil fractions boiling between about 400 and 800 F. and of high naphthenes content to produce motor fuel of high octane number which comprises feeding said fractions in contact with a finely divided fluidized solid catalyst containing 84 ,to 92% by weight of alumina-silica cracking catalyst containing a major proportion of silica and 8 to 16% of molybdenum oxide, said operations being conducted at a pressure of 150 to 300 p.s.i.g. at a temperature of 1000 to 1100 F. in the presence of hydrogen supplied at the rate of 3000 to 10,000 cubic feet per barrel of feed.

3. Process according to claim 2 wherein the operating temperature is between 1035 and 1070 F.

4. Process according to claim 2 wherein the pressure is about 200 p.s.i.g.

5. A process according to claim 2 wherein the proportion of molybdenum oxide is about 10% of the total catalyst.

6. The process of cracking petroleum oil which comprises feeding a gas oil of high naphthenes content and boiling range between 400 and 800 F. to a hydrocracking reactor in the presence of finely divided solid catalyst comprising an alumina-silica cracking catalyst containing a major proportion of silica containing 8 to 16% by weight of molybdenum oxide, said reactor being operated at a pressure of about 150-300 p.s.i.g. and a temperature of 1035 to 1070 F., feeding hydrogen gas along with said oil and catalyst at a rate of about 4,000 cubic feet per barrel of gas oil, simultaneously subjecting a naphtha to a hydroforming operation in the presence of a hydrogenation catalyst, and utilizing the hydrogen from the hydroforming operation to supply hydrogen for the hydrocracking operation.

7. The process which comprises hydrocracking a gas oil over a silica-alumina base cracking catalyst containing a major proportion of silica which contains 8 to 16% by weight of molybdenum oxide, to produce liquid hydrocarbons of the motor fuel boiilng range and of octane number at least plus gaseous hydrocarbons in which molybdenum oxide is entrained, conveying the gases with entrained molybdenum oxide to a zone containing fluidized base catalyst and there reimpregnating the base catalyst with said entrained moylbdenum oxide, and returning the reimpregnated catalyst to the hydrocracking step.

References Cited in the file of this patent UNITED STATES PATENTS Thomas July 1, 1947 (Other references on following page) 13 UNITED STATES PATENTS Hulfman Mar. 9, 1948 Munday June 14, 1949 McKinley Feb. 12, 1950 Parker Feb. 28, 1950 Wilson Feb. 13, 1951 Haensel Dec. 29, 1953 14 2,700,014 Anhorn et al. Jan. 18, 1955 2,703,308 Oblad et a1. Mar. 1, 1955 2,799,626 Johnson et a1. July 16, 1957 OTHER REFERENCES Control of Physical Structure of Silica-Alumina Catalyst, Ashley et 211., Industrial and Eng. Chem., vol. 44, No. 12, December 1952, page 2860, Table III. 

1. A PROCESS OF CRACKING MINERAL OIL FRACTIONS BOILING ABOVE THE GASOLINE RANGE TO PRODUCE MOTOR FUEL OF HIGH OCTANE NUMBER WITH LOW CARBON PRODUCTION WHICH COMPRISES SUBJECTING SAID FRACTIONS TO CONTACT WITH A CRACKING CATALYST CONTAINING A MAJOR PROPORTION OF SILICA, A MINOR PROPORTION OF A REFRACTORY METAL OXIDE SELECTED FROM THE GROUP CONSISTING OF ALUMINA AND MAGNESIA AND FROM 8 TO 16% OF MOLYBDENUM OXIDE AT A TEMPERATURE OF 1000 TO 1100*F. AND A PRESSURE OF 150 TO 300 POUNDS'' PRESSURE PER SQUARE INCH AND IN THE PRESENCE OF HYDROGEN GAS SUPPLIED AT THE RATE OF 3000 TO 10,000 CUBIC FEET UNDER STANDARD CONDITIONS PER BARREL OF FEED.
 6. THE PROCESS OF CRACKING PETROLEUM OIL WHICH COMPRIESE FEEDING A GAS OIL HIGH NAPHTHENES CONTENT AND BOILING RANGE BETWEEN 400 AND 800*F. TO A HYDROCRACKING REACTOR IN THE PRESENCE OF FINELY DIVIDED SOLID CATALYST COMPRISING AN ALUMINA-SILICA CRACKING CATALYST CONTAINING A MAJOR PROPORTION OF SILICA CONTAINING 8 TO 16% BY WEIGHT OF MOLYBDENUM OXIDE, SAID REACTOR BEING OPERATED AT A PRESSURE OF ABOUT 150-300 P.S.I.G. AND A TEMPERATURE OF 1035* TO 1070*F., FEEDING HYDROGEN GAS ALONG WITH SAID OIL AND CATALYST AT A RATE OF ABOUT 4,000 CUBIC FEED PER BARREL OF GAS OIL, SIMULTANEOUSLY SUBJECTING A NAPHTHA TO A HYDROFORMING OPERATION IN THE PRESENCE OF A HYDROGENATION CATALYST, AND UTILIZING THE HYDROGEN FROM THE HYDROFORMING OPERATION TO SUPPLY HYDROGEN FOR THE HYDROCRACKING OPERATION. 